This invention relates to a process for the separation of ethane and heavier hydrocarbons or propane and heavier hydrocarbons from liquefied natural gas (hereinafter referred to as LNG) combined with the separation of a gas containing hydrocarbons to provide a volatile methane-rich gas stream and a less volatile natural gas liquids (NGL) or liquefied petroleum gas (LPG) stream.
As an alternative to transportation in pipelines, natural gas at remote locations is sometimes liquefied and transported in special LNG tankers to appropriate LNG receiving and storage terminals. The LNG can then be re-vaporized and used as a gaseous fuel in the same fashion as natural gas. Although LNG usually has a major proportion of methane, i.e., methane comprises at least 50 mole percent of the LNG, it also contains relatively lesser amounts of heavier hydrocarbons such as ethane, propane, butanes, and the like, as well as nitrogen. It is often necessary to separate some or all of the heavier hydrocarbons from the methane in the LNG so that the gaseous fuel resulting from vaporizing the LNG conforms to pipeline specifications for heating value. In addition, it is often also desirable to separate the heavier hydrocarbons from the methane and ethane because these hydrocarbons have a higher value as liquid products (for use as petrochemical feedstocks, as an example) than their value as fuel.
Although there are many processes which may be used to separate ethane and/or propane and heavier hydrocarbons from LNG, these processes often must compromise between high recovery, low utility costs, and process simplicity (and hence low capital investment). U.S. Pat. Nos. 2,952,984; 3,837,172; 5,114,451; and 7,155,931 describe relevant LNG processes capable of ethane or propane recovery while producing the lean LNG as a vapor stream that is thereafter compressed to delivery pressure to enter a gas distribution network. However, lower utility costs may be possible if the lean LNG is instead produced as a liquid stream that can be pumped (rather than compressed) to the delivery pressure of the gas distribution network, with the lean LNG subsequently vaporized using a low level source of external heat or other means. U.S. Pat. Nos. 6,604,380; 6,907,752; 6,941,771; 7,069,743; and 7,216,507 and co-pending application Ser. Nos. 11/749,268 and 12/060,362 describe such processes.
Economics and logistics often dictate that LNG receiving terminals be located close to the natural gas transmission lines that will transport the re-vaporized LNG to consumers. In many cases, these areas also have plants for processing natural gas produced in the region to recover the heavier hydrocarbons contained in the natural gas. Available processes for separating these heavier hydrocarbons include those based upon cooling and refrigeration of gas, oil absorption, and refrigerated oil absorption. Additionally, cryogenic processes have become popular because of the availability of economical equipment that produces power while simultaneously expanding and extracting heat from the gas being processed. Depending upon the pressure of the gas source, the richness (ethane, ethylene, and heavier hydrocarbons content) of the gas, and the desired end products, each of these processes or a combination thereof may be employed.
The cryogenic expansion process is now generally preferred for natural gas liquids recovery because it provides maximum simplicity with ease of startup, operating flexibility, good efficiency, safety, and good reliability. U.S. Pat. Nos. 3,292,380; 4,061,481; 4,140,504; 4,157,904; 4,171,964; 4,185,978; 4,251,249; 4,278,457; 4,519,824; 4,617,039; 4,687,499; 4,689,063; 4,690,702; 4,854,955; 4,869,740; 4,889,545; 5,275,005; 5,555,748; 5,566,554; 5,568,737; 5,771,712; 5,799,507; 5,881,569; 5,890,378; 5,983,664; 6,182,469; 6,578,379; 6,712,880; 6,915,662; 7,191,617; 7,219,513; reissue U.S. Pat. No. 33,408; and co-pending application Ser. Nos. 11/430,412; 11/839,693; 11/971,491; and 12/206,230 describe relevant processes (although the description of the present invention is based on different processing conditions than those described in the cited U.S. patents).
The present invention is generally concerned with the integrated recovery of ethylene, ethane, propylene, propane, and heavier hydrocarbons from such LNG and gas streams. It uses a novel process arrangement to integrate the heating of the LNG stream and the cooling of the gas stream to eliminate the need for a separate vaporizer and the need for external refrigeration, allowing high C2 component recovery while keeping the processing equipment simple and the capital investment low. Further, the present invention offers a reduction in the utilities (power and heat) required to process the LNG and gas streams, resulting in lower operating costs than other processes, and also offering significant reduction in capital investment.
Heretofore, assignee's U.S. Pat. No. 7,216,507 has been used to recover C2 components and heavier hydrocarbon components in plants processing LNG, while assignee's co-pending application Ser. No. 11/430,412 could be used to recover C2 components and heavier hydrocarbon components in plants processing natural gas. Surprisingly, applicants have found that by integrating certain features of the assignee's U.S. Pat. No. 7,216,507 invention with certain features of the assignee's co-pending application Ser. No. 11/430,412, extremely high C2 component recovery levels can be accomplished using less energy than that required by individual plants to process the LNG and natural gas separately.
A typical analysis of an LNG stream to be processed in accordance with this invention would be, in approximate mole percent, 92.2% methane, 6.0% ethane and other C2 components, 1.1% propane and other C3 components, and traces of butanes plus, with the balance made up of nitrogen. A typical analysis of a gas stream to be processed in accordance with this invention would be, in approximate mole percent, 80.1% methane, 9.5% ethane and other C2 components, 5.6% propane and other C3 components, 1.3% iso-butane, 1.1% normal butane, 0.8% pentanes plus, with the balance made up of nitrogen and carbon dioxide. Sulfur containing gases are also sometimes present.
FIGS. 1 and 2 are provided to quantify the advantages of the present invention.
In the following explanation of the above figures, tables are provided summarizing flow rates calculated for representative process conditions. In the tables appearing herein, the values for flow rates (in moles per hour) have been rounded to the nearest whole number for convenience. The total stream rates shown in the tables include all non-hydrocarbon components and hence are generally larger than the sum of the stream flow rates for the hydrocarbon components. Temperatures indicated are approximate values rounded to the nearest degree. It should also be noted that the process design calculations performed for the purpose of comparing the processes depicted in the figures are based on the assumption of no heat leak from (or to) the surroundings to (or from) the process. The quality of commercially available insulating materials makes this a very reasonable assumption and one that is typically made by those skilled in the art.
For convenience, process parameters are reported in both the traditional British units and in the units of the Système International d'Unités (SI). The molar flow rates given in the tables may be interpreted as either pound moles per hour or kilogram moles per hour. The energy consumptions reported as horsepower (HP) and/or thousand British Thermal Units per hour (MBTU/Hr) correspond to the stated molar flow rates in pound moles per hour. The energy consumptions reported as kilowatts (kW) correspond to the stated molar flow rates in kilogram moles per hour.
FIG. 1 is a flow diagram showing the design of a processing plant to recover C2+ components from natural gas using an LNG stream to provide refrigeration. In the simulation of the FIG. 1 process, inlet gas enters the plant at 126° F. [52° C.] and 600 psia [4,137 kPa(a)] as stream 31. If the inlet gas contains a concentration of sulfur compounds which would prevent the product streams from meeting specifications, the sulfur compounds are removed by appropriate pretreatment of the feed gas (not illustrated). In addition, the feed stream is usually dehydrated to prevent hydrate (ice) formation under cryogenic conditions. Solid desiccant has typically been used for this purpose.
The inlet gas stream 31 is cooled in heat exchanger 12 by heat exchange with a portion (stream 72a) of partially warmed LNG at −174° F. [−114° C.] and cool distillation stream 38a at −107° F. [−77° C.]. The cooled stream 31a enters separator 13 at −79° F. [−62° C.] and 584 psia [4,027 kPa(a)] where the vapor (stream 34) is separated from the condensed liquid (stream 35). Liquid stream 35 is flash expanded through an appropriate expansion device, such as expansion valve 17, to the operating pressure (approximately 430 psia [2,965 kPa(a)]) of fractionation tower 20. The expanded stream 35a leaving expansion valve 17 reaches a temperature of −93° F. [−70° C.] and is supplied to fractionation tower 20 at a first mid-column feed point.
The vapor from separator 13 (stream 34) enters a work expansion machine 10 in which mechanical energy is extracted from this portion of the high pressure feed. The machine 10 expands the vapor substantially isentropically to slightly above the tower operating pressure, with the work expansion cooling the expanded stream 34a to a temperature of approximately −101° F. [−74° C.]. The typical commercially available expanders are capable of recovering on the order of 80-88% of the work theoretically available in an ideal isentropic expansion. The work recovered is often used to drive a centrifugal compressor (such as item 11) that can be used to re-compress the heated distillation stream (stream 38b), for example. The expanded stream 34a is further cooled to −124° F. [−87° C.] in heat exchanger 14 by heat exchange with cold distillation stream 38 at −143° F. [−97° C.], whereupon the partially condensed expanded stream 34b is thereafter supplied to fractionation tower 20 at a second mid-column feed point.
The demethanizer in tower 20 is a conventional distillation column containing a plurality of vertically spaced trays, one or more packed beds, or some combination of trays and packing to provide the necessary contact between the liquids falling downward and the vapors rising upward. The column also includes reboilers (such as reboiler 19) which heat and vaporize a portion of the liquids flowing down the column to provide the stripping vapors which flow up the column to strip the liquid product, stream 41, of methane and lighter components. Liquid product stream 41 exits the bottom of the tower at 99° F. [37° C.], based on a typical specification of a methane to ethane ratio of 0.020:1 on a molar basis in the bottom product.
Overhead distillation stream 43 is withdrawn from the upper section of fractionation tower 20 at −143° F. [−97° C.] and is divided into two portions, streams 44 and 47. The first portion, stream 44, flows to reflux condenser 23 where it is cooled to −237° F. [−149° C.] and totally condensed by heat exchange with a portion (stream 72) of the cold LNG (stream 71a). Condensed stream 44a enters reflux separator 24 wherein the condensed liquid (stream 46) is separated from any uncondensed vapor (stream 45). The liquid stream 46 from reflux separator 24 is pumped by reflux pump 25 to a pressure slightly above the operating pressure of demethanizer 20 and stream 46a is then supplied as cold top column feed (reflux) to demethanizer 20. This cold liquid reflux absorbs and condenses the C2 components and heavier hydrocarbon components from the vapors rising in the upper section of demethanizer 20.
The second portion (stream 47) of overhead vapor stream 43 combines with any uncondensed vapor (stream 45) from reflux separator 24 to form cold distillation stream 38 at −143° F. [−97° C.]. Distillation stream 38 passes countercurrently to expanded stream 34a in heat exchanger 14 where it is heated to −107° F. [−77° C.] (stream 38a), and countercurrently to inlet gas in heat exchanger 12 where it is heated to 47° F. [8° C.] (stream 38b). The distillation stream is then re-compressed in two stages. The first stage is compressor 11 driven by expansion machine 10. The second stage is compressor 21 driven by a supplemental power source which compresses stream 38c to sales line pressure (stream 38d). After cooling to 126° F. [52° C.] in discharge cooler 22, stream 38e combines with warm LNG stream 71b to form the residue gas product (stream 42). Residue gas stream 42 flows to the sales gas pipeline at 1262 psia [8,701 kPa(a)], sufficient to meet line requirements.
The LNG (stream 71) from LNG tank 50 enters pump 51 at −251° F. [−157° C.]. Pump 51 elevates the pressure of the LNG sufficiently so that it can flow through heat exchangers and thence to the sales gas pipeline. Stream 71a exits the pump 51 at −242° F. [−152° C.] and 1364 psia [9,404 kPa(a)] and is divided into two portions, streams 72 and 73. The first portion, stream 72, is heated as described previously to −174° F. [−114° C.] in reflux condenser 23 as it provides cooling to the portion (stream 44) of overhead vapor stream 43 from fractionation tower 20, and to 43° F. [6° C.] in heat exchanger 12 as it provides cooling to the inlet gas. The second portion, stream 73, is heated to 35° F. [2° C.] in heat exchanger 53 using low level utility heat. The heated streams 72b and 73a recombine to form warm LNG stream 71b at 40° F. [4° C.], which thereafter combines with distillation stream 38e to form residue gas stream 42 as described previously.
A summary of stream flow rates and energy consumption for the process illustrated in FIG. 1 is set forth in the following table:
TABLE I(FIG. 1)Stream Flow Summary - Lb. Moles/Hr [kg moles/Hr]StreamMethaneEthanePropaneButanes+Total3142,5455,0482,9721,65853,1453433,4811,6062793936,221359,0643,4422,6931,61916,9244350,499250051,534448,0554008,2214500000468,0554008,2214742,444210043,3133842,444210043,3137140,2932,642491343,6897227,6011,810336229,9277312,692832155113,7624282,7372,663491387,002411015,0272,9721,6589,832Recoveries*Ethane65.37%Propane85.83%Butanes+99.83%PowerLNG Feed Pump3,561HP[5,854kW]Reflux Pump23HP[38kW]Residue Gas Compressor24,612HP[40,462kW]Totals28,196HP[46,354kW]Low Level Utility HeatLNG Heater68,990MBTU/Hr[44,564kW]High Level Utility HeatDemethanizer Reboiler80,020MBTU/Hr[51,689kW]Specific PowerHP-Hr/Lb. Mole2.868[kW-Hr/kg mole][4.715]*(Based on un-rounded flow rates)
The recoveries reported in Table I are computed relative to the total quantities of ethane, propane, and butanes+ contained in the gas stream being processed in the plant and in the LNG stream. Although the recoveries are quite high relative to the heavier hydrocarbons contained in the gas being processed (99.58%, 100.00%, and 100.00%, respectively, for ethane, propane, and butanes+), none of the heavier hydrocarbons contained in the LNG stream are captured in the FIG. 1 process. In fact, depending on the composition of LNG stream 71, the residue gas stream 42 produced by the FIG. 1 process may not meet all pipeline specifications. The specific power reported in Table I is the power consumed per unit of liquid product recovered, and is an indicator of the overall process efficiency.
FIG. 2 is a flow diagram showing processes to recover C2+ components from LNG and natural gas in accordance with U.S. Pat. No. 7,216,507 and co-pending application Ser. No. 11/430,412, respectively, with the processed LNG stream used to provide refrigeration for the natural gas plant. The processes of FIG. 2 have been applied to the same LNG stream and inlet gas stream compositions and conditions as described previously for FIG. 1.
In the simulation of the FIG. 2 process, the LNG to be processed (stream 71) from LNG tank 50 enters pump 51 at −251° F. [−157° C.]. Pump 51 elevates the pressure of the LNG sufficiently so that it can flow through heat exchangers and thence to expansion machine 55. Stream 71a exits the pump at −242° F. [−152° C.] and 1364 psia [9,404 kPa(a)] and is split into two portions, streams 75 and 76. The first portion, stream 75, is expanded to the operating pressure (approximately 415 psia [2,859 kPa(a)]) of fractionation column 62 by expansion valve 58. The expanded stream 75a leaves expansion valve 58 at −238° F. [−150° C.] and is thereafter supplied to tower 62 at an upper mid-column feed point.
The second portion, stream 76, is heated to −79° F. [−62° C.] in heat exchanger 52 by cooling compressed overhead distillation stream 79a at −70° F. [−57° C.] and reflux stream 82 at −128° F. [−89° C.]. The partially heated stream 76a is further heated and vaporized in heat exchanger 53 using low level utility heat. The heated stream 76b at −5° F. [−20° C.] and 1334 psia [9,198 kPa(a)] enters work expansion machine 55 in which mechanical energy is extracted from this portion of the high pressure feed. The machine 55 expands the vapor substantially isentropically to the tower operating pressure, with the work expansion cooling the expanded stream 76c to a temperature of approximately −107° F. [−77° C.] before it is supplied as feed to fractionation column 62 at a lower mid-column feed point.
The demethanizer in fractionation column 62 is a conventional distillation column containing a plurality of vertically spaced trays, one or more packed beds, or some combination of trays and packing consisting of two sections. The upper absorbing (rectification) section contains the trays and/or packing to provide the necessary contact between the vapors rising upward and cold liquid falling downward to condense and absorb the ethane and heavier components; the lower stripping (demethanizing) section contains the trays and/or packing to provide the necessary contact between the liquids falling downward and the vapors rising upward. The demethanizing section also includes one or more reboilers (such as side reboiler 60 using low level utility heat, and reboiler 61 using high level utility heat) which heat and vaporize a portion of the liquids flowing down the column to provide the stripping vapors which flow up the column. The column liquid stream 80 exits the bottom of the tower at 54° F. [12° C.], based on a typical specification of a methane to ethane ratio of 0.020:1 on a molar basis in the bottom product.
Overhead distillation stream 79 is withdrawn from the upper section of fractionation tower 62 at −144° F. [−98° C.] and flows to compressor 56 driven by expansion machine 55, where it is compressed to 807 psia [5,567 kPa(a)] (stream 79a). At this pressure, the stream is totally condensed as it is cooled to −128° F. [−89° C.] in heat exchanger 52 as described previously. The condensed liquid (stream 79b) is then divided into two portions, streams 83 and 82. The first portion (stream 83) is the methane-rich lean LNG stream, which is pumped by pump 63 to 1278 psia [8,809 kPa(a)] for subsequent vaporization in heat exchangers 14 and 12, heating stream 83a to −114° F. [−81° C.] and then to 40° F. [4° C.] as described in paragraphs [0036] and [0033] below to produce warm lean LNG stream 83c. 
The remaining portion of condensed liquid stream 79b, reflux stream 82, flows to heat exchanger 52 where it is subcooled to −237° F. [−149° C.] by heat exchange with a portion of the cold LNG (stream 76) as described previously. The subcooled stream 82a is then expanded to the operating pressure of demethanizer 62 by expansion valve 57. The expanded stream 82b at −236° F. [−149° C.] is then supplied as cold top column feed (reflux) to demethanizer 62. This cold liquid reflux absorbs and condenses the C2 components and heavier hydrocarbon components from the vapors rising in the upper rectification section of demethanizer 62.
In the simulation of the FIG. 2 process, inlet gas enters the plant at 126° F. [52° C.] and 600 psia [4,137 kPa(a)] as stream 31. The feed stream 31 is cooled in heat exchanger 12 by heat exchange with cool lean LNG (stream 83b), cool overhead distillation stream 38a at −114° F. [−81° C.], and demethanizer liquids (stream 39) at −51° F. [−46° C.]. The cooled stream 31a enters separator 13 at −91° F. [−68° C.] and 584 psia [4,027 kPa(a)] where the vapor (stream 34) is separated from the condensed liquid (stream 35). Liquid stream 35 is flash expanded through an appropriate expansion device, such as expansion valve 17, to the operating pressure (approximately 390 psia [2,687 kPa(a)]) of fractionation tower 20. The expanded stream 35a leaving expansion valve 17 reaches a temperature of −111° F. [−80° C.] and is supplied to fractionation tower 20 at a first lower mid-column feed point.
Vapor stream 34 from separator 13 enters a work expansion machine 10 in which mechanical energy is extracted from this portion of the high pressure feed. The machine 10 expands the vapor substantially isentropically to the tower operating pressure, with the work expansion cooling the expanded stream 34a to a temperature of approximately −121° F. [−85° C.]. The partially condensed expanded stream 34a is thereafter supplied as feed to fractionation tower 20 at a second lower mid-column feed point.
The demethanizer in fractionation column 20 is a conventional distillation column containing a plurality of vertically spaced trays, one or more packed beds, or some combination of trays and packing consisting of two sections. The upper absorbing (rectification) section contains the trays and/or packing to provide the necessary contact between the vapors rising upward and cold liquid falling downward to condense and absorb the ethane and heavier components; the lower stripping (demethanizing) section contains the trays and/or packing to provide the necessary contact between the liquids falling downward and the vapors rising upward. The demethanizing section also includes one or more reboilers (such as the side reboiler in heat exchanger 12 described previously, and reboiler 19 using high level utility heat) which heat and vaporize a portion of the liquids flowing down the column to provide the stripping vapors which flow up the column. The column liquid stream 40 exits the bottom of the tower at 89° F. [31° C.], based on a typical specification of a methane to ethane ratio of 0.020:1 on a molar basis in the bottom product, and combines with stream 80 to form the liquid product (stream 41).
A portion of the distillation vapor (stream 44) is withdrawn from the upper region of the stripping section of fractionation column 20 at −125° F. [−87° C.] and compressed to 545 psia [3,756 kPa(a)] by compressor 26. The compressed stream 44a is then cooled from −87° F. [−66° C.] to −143° F. [−97° C.] and condensed (stream 44b) in heat exchanger 14 by heat exchange with cold overhead distillation stream 38 exiting the top of demethanizer 20 and cold lean LNG (stream 83a) at −116° F. [−82° C.]. Condensed liquid stream 44b is expanded by expansion valve 16 to a pressure slightly above the operating pressure of demethanizer 20, and the resulting stream 44c at −146° F. [−99° C.] is then supplied as cold liquid reflux to an intermediate region in the absorbing section of demethanizer 20. This supplemental reflux absorbs and condenses most of the C3 components and heavier components (as well as some of the C2 components) from the vapors rising in the lower rectification region of the absorbing section so that only a small amount of recycle (stream 36) must be cooled, condensed, subcooled, and flash expanded to produce the top reflux stream 36c that provides the final rectification in the upper region of the absorbing section of demethanizer 20. As the cold reflux stream 36c contacts the rising vapors in the upper region of the absorbing section, it condenses and absorbs the C2 components and any remaining C3 components and heavier components from the vapors so that they can be captured in the bottom product (stream 40) from demethanizer 20.
Overhead distillation stream 38 is withdrawn from the upper section of fractionation tower 20 at −148° F. [−100° C.]. It passes countercurrently to compressed distillation vapor stream 44a and recycle stream 36a in heat exchanger 14 where it is heated to −114° F. [−81° C.] (stream 38a), and countercurrently to inlet gas stream 31 and recycle stream 36 in heat exchanger 12 where it is heated to 20° F. [−7° C.] (stream 38b). The distillation stream is then re-compressed in two stages. The first stage is compressor 11 driven by expansion machine 10. The second stage is compressor 21 driven by a supplemental power source which compresses stream 38c to sales line pressure (stream 38d). After cooling to 126° F. [52° C.] in discharge cooler 22, stream 38e is divided into two portions, stream 37 and recycle stream 36. Stream 37 combines with warm lean LNG stream 83c to form the residue gas product (stream 42). Residue gas stream 42 flows to the sales gas pipeline at 1262 psia [8,701 kPa(a)], sufficient to meet line requirements.
Recycle stream 36 flows to heat exchanger 12 and is cooled to −105° F. [−76° C.] by heat exchange with cool lean LNG (stream 83b), cool overhead distillation stream 38a, and demethanizer liquids (stream 39) as described previously. Stream 36a is further cooled to −143° F. [−97° C.] by heat exchange with cold lean LNG stream 83a and cold overhead distillation stream 38 in heat exchanger 14 as described previously. The substantially condensed stream 36b is then expanded through an appropriate expansion device, such as expansion valve 15, to the demethanizer operating pressure, resulting in cooling of the total stream to −151° F. [−102° C.]. The expanded stream 36c is then supplied to fractionation tower 20 as the top column feed. Any vapor portion of stream 36c combines with the vapors rising from the top fractionation stage of the column to form overhead distillation stream 38, which is withdrawn from an upper region of the tower as described previously.
A summary of stream flow rates and energy consumption for the process illustrated in FIG. 2 is set forth in the following table:
TABLE II(FIG. 2)Stream Flow Summary - Lb. Moles/Hr [kg moles/Hr]StreamMethaneEthanePropaneButanes+Total3142,5455,0482,9721,65853,1453428,7621,0511632230,7593513,7833,9972,8091,63622,386446,746195307,0003849,040390050,064366,5955006,7333742,445340043,331401005,0142,9721,6589,8147140,2932,642491343,689754,8353175905,2437635,4582,325432338,4467945,588160045,898825,3482005,3858340,240140040,51380532,62849133,1764282,685480083,844411537,6423,4631,66112,990Recoveries*Ethane 99.38%Propane100.00%Butanes+100.00%PowerLNG Feed Pump3,552HP[5,839kW]LNG Product Pump1,774HP[2,916kW]Residue Gas Compressor29,272HP[48,123kW]Reflux Compressor601HP[988kW]Totals35,199HP[57,866kW]Low Level Utility HeatLiquid Feed Heater66,200MBTU/Hr[42,762kW]Demethanizer Reboiler 6023,350MBTU/Hr[15,083kW]Totals89,550MBTU/Hr[57,845kW]High Level Utility HeatDemethanizer Reboiler 1926,780MBTU/Hr[17,298kW]Demethanizer Reboiler 613,400MBTU/Hr[2,196kW]Totals30,180MBTU/Hr[19,494kW]Specific PowerHP-Hr/Lb. Mole2.710[kW-Hr/kg mole][4.455]*(Based on un-rounded flow rates)
Comparison of the recovery levels displayed in Tables I and II shows that the liquids recovery of the FIG. 2 processes is much higher than that of the FIG. 1 process due to the recovery of the heavier hydrocarbon liquids contained in the LNG stream in fractionation tower 62. The ethane recovery improves from 65.37% to 99.38%, the propane recovery improves from 85.83% to 100.00%, and the butanes+recovery improves from 99.83% to 100.00%. In addition, the process efficiency of the FIG. 2 processes is improved by more than 5% in terms of the specific power relative to the FIG. 1 process.